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| United States Patent |
5,043,525
|
|
Haizmann
, et al.
|
August 27, 1991
|
Paraffin isomerization and liquid phase adsorptive product separation
Abstract
A combination isomerization and liquid phase adsorptive separation process
is given increased efficiency and cost effectiveness while also improving
the product quality by eliminating the columns for the separation of
desorbent material from extract and raffinate streams. In this arrangement
a C.sub.5 + naphtha stream is split into a heavy hydrocarbon stream
comprising normal hexane and higher boiling hydrocarbons and an
isomerization zone feedstream comprising isohexane and lower boiling
hydrocarbons. The heavy hydrocarbon stream goes directly to a
deisohexanizer column. The isomerization zone feedstream is combined with
an excess desorbent stream and the extract stream from an adsorptive
separation section to form a combined feed. Hydrocarbons in the combined
feed are isomerized and after stabilization pass directly into the
adsorption section. In the adsorption section, normal pentanes are
selectively adsorbed on an adsorbent material, and a raffinate stream
comprising desorbent and isoparaffins is passed to the deisohexanizer
column and supplies the desorbent for the adsorption section. Any
desorbent in excess of that required for the adsorption section is
combined with the isomerization zone feed. The extract stream that is
combined with the isomerization zone feed is recovered from adsorption
section. A bottoms stream comprising C.sub.7 and higher boiling
hydrocarbons is withdrawn from the bottom of the deisohexanizer column. A
high octane isomerate is taken overhead from the deisohexanizer as a
product stream.
| Inventors:
|
Haizmann; Robert S. (Rolling Meadows, IL);
Hibbs; Frederick M. (London, GB2);
Raghuram; Srikantiah (Darien, IL)
|
| Assignee:
|
UOP (Des Plaines, IL)
|
| Appl. No.:
|
560655 |
| Filed:
|
July 30, 1990 |
| Current U.S. Class: |
585/737; 585/738; 585/748; 585/821; 585/826 |
| Intern'l Class: |
C07C 005/13; C07C 007/12 |
| Field of Search: |
585/737,738,748,821,826
|
References Cited
U.S. Patent Documents
| 2966528 | Dec., 1960 | Haensel | 260/666.
|
| 3755144 | Aug., 1973 | Asselin | 208/95.
|
| 4210771 | Jul., 1980 | Holcombe | 585/738.
|
| 4717784 | Jan., 1988 | Stem et al. | 585/738.
|
| 4804802 | Feb., 1989 | Evans et al. | 585/734.
|
| 4956521 | Sep., 1990 | Volles | 585/737.
|
Primary Examiner: McFarlane; Anthony
Assistant Examiner: Phan; Nhat
Attorney, Agent or Firm: McBride; Thomas K., Tolomei; John G.
Claims
What is claimed is:
1. A process for the isomerization of a feedstream comprising C.sub.5
-C.sub.6 hydrocarbons said process comprising:
(a) separating said feedstream comprising C.sub.5 -C.sub.6 hydrocarbons
into a first stream comprising normal hexane and higher boiling
hydrocarbons and a second stream comprising components from said
feedstream boiling below normal hexane;
(b) combining said second stream with an extract stream to produce a
combined feed and passing said combined feed into an isomerization zone
containing an isomerization catalyst at isomerization conditions to
isomerize the hydrocarbons in said second stream and recovering an
isomerization zone effluent from said isomerization zone comprising
C.sub.5 and C.sub.6 isoparaffins;
(c) passing at least a portion of the effluent from said isomerization zone
to an adsorption section, contacting said effluent with an adsorbent,
selectively adsorbing normal paraffins from said effluent, and producing a
raffinate stream having a decreased concentration of at least one normal
paraffin component relative to said effluent stream and a adsorbent
material having normal paraffins adsorbed thereon;
d) passing at least a portion of said raffinate stream and said first
stream into a fractionation zone, withdrawing hydrocarbons having boiling
points greater than normal hexane from the bottom of said fractionation
zone, recovering an isomerate product from an overhead of said
fractionation zone and withdrawing a sidecut stream comprising normal
hexane from said fractionation zone as a desorbent stream;
(e) passing said desorbent stream to said adsorption zone and contacting
said adsorbent material having normal paraffins adsorbed thereon with said
desorbent stream to desorb at least one normal paraffin component and
produce said extract stream.
2. The process of claim 1 wherein said feedstream includes C7 and higher
boiling hydrocarbons.
3. The process of claim 1 wherein said combined stream is formed by
combining said extract stream and said second stream with a portion of
said desorbent stream.
4. The process of claim 1 wherein said adsorption zone comprises a
simulated moving bed adsorption zone.
5. The process of claim 1 wherein said isomerization effluent passes
directly to a stabilizer, C.sub.4 and lighter hydrocarbons are removed
from said effluent and the remainder of the effluent is passed directly to
said adsorption zone.
6. The process of claim 1 wherein the adsorbent in said adsorption zone
comprises calcium zeolite A.
7. The process of claim 1 wherein said adsorption zone operates in liquid
phase.
8. The process of claim 1 wherein said isomerization catalyst comprises
alumina having from 0.01 to 0.25 wt. % platinum and from 2 to 10 wt. % of
a chloride component.
9. The process of claim 4 wherein said adsorption section comprises at
least one bed of adsorbent, the total quantity of adsorbent is divided
into at least three operationally distinct zones of adsorbent, said
effluent and desorbent are charged to different zones and the position of
the zones relative to said total quantity of said adsorbent is at least
intermittently varied by changing withdrawal and input points for said
feed and desorbent.
10. The process of claim 1 wherein said fractionation zone comprises a
single column, said first stream enters said column at a first inlet point
lower elevation, said raffinate stream enters said column at a second
inlet point located above said first inlet point and said sidecut is
withdrawn from said column at an elevation between said first and second
inlet points.
11. The process of claim 1 wherein said raffinate stream has a decreased
concentration of normal pentane relative to said effluent stream.
12. The process of claim 1 wherein said desorbent stream comprises normal
hexane and mono-methyl pentanes.
13. A process for the isomerization of a C.sub.5 +naphtha feedstream
comprising C.sub.5 -C.sub.6 hydrocarbons said process comprising:
(a) separating said feedstream in a splitter column into a first stream
comprising normal hexane and higher boiling hydrocarbons and a second
stream comprising components from said feedstream boiling below normal
hexane;
(b) combining said second stream with an extract stream and a portion of a
desorbent stream to produce a combined feed and passing said combined feed
into an isomerization zone containing an isomerization catalyst and
contacting said combined feed at isomerization conditions to isomerize the
hydrocarbons in said second stream and recovering an isomerization zone
effluent from said isomerization zone comprising C.sub.5 and C.sub.6
isoparaffins;
(c) separating C.sub.4 and lower boiling hydrocarbons from said effluent
stream;
(d) passing at least a portion of the effluent from said isomerization zone
to a simulated moving bed adsorption zone, contacting said effluent with
an adsorbent, selectively adsorbing normal paraffins from said effluent,
and producing a raffinate stream having a reduced concentration of normal
pentane relative to said effluent stream and an adsorbent material having
normal pentane adsorbed thereon;
(e) passing at least a portion of said raffinate stream into an upper
portion of a fractionation column, passing at least a portion of said
first stream into a lower half of said fractionation column and
withdrawing feedstream components having a boiling point greater than
normal hexane from the bottom of said fractionation column, recovering an
isomerate product stream from the top of said fractionation column, said
isomerate product stream being essentially free of normal hexane and
higher boiling hydrocarbons, and withdrawing said desorbent stream as a
sidecut from said column at a tray location intermediate the column
locations where said first stream and said raffinate stream enter said
column, said desorbent stream comprising normal hexane and lower boiling
hydrocarbons; and,
(f) passing at least a portion of said desorbent stream to said adsorption
zone and contacting said adsorbent material having normal pentane adsorbed
thereon with said desorbent stream to desorb said normal pentane and
produce said extract stream.
Description
BACKGROUND OF THE INVENTION
This invention relates generally to the isomerization of hydrocarbons. This
invention relates more specifically to the isomerization of light
paraffins using a solid catalyst, and the separation of more highly
branched paraffins from less highly branched paraffins by adsorptive
separation.
DESCRIPTION OF THE PRIOR ART
High octane gasoline is required for modern gasoline engines. Formerly it
was common to accomplish octane number improvement by the use of various
lead-containing additives. As lead is phased out of gasoline for
environmental reasons, it has become increasingly necessary to rearrange
the structure of the hydrocarbons used in gasoline blending in order to
obtain high octane levels. Catalytic reforming and catalytic isomerization
are two widely used processes for this upgrading.
A gasoline blending pool normally includes C.sub.4 and heavier hydrocarbons
having boiling points of less than 205.degree. C. (395.degree. F.) at
atmospheric pressure. This range of hydrocarbon includes C.sub.4 -C.sub.6
paraffins and especially the C.sub.5 and C.sub.6 normal paraffins which
have relatively low octane numbers. The C.sub.4 -C.sub.6 hydrocarbons have
the greatest susceptibility to octane improvement by lead addition and
were formerly upgrade in this manner. Octane improvement can also be
obtained by using isomerization to rearrange the structure of the
paraffinic hydrocarbons into branched-chain paraffins or reforming to
convert the C.sub.6 and heavier hydrocarbons to aromatic compounds. Normal
C.sub.5 hydrocarbons are not readily converted into aromatics, therefore,
the common practice has been to isomerize these lighter hydrocarbons into
corresponding branched-chain isoparaffins. Although the C.sub.6 and
heavier hydrocarbons can be upgraded into aromatics through
hydrocyclization, the conversion of C.sub.6 's to aromatics creates higher
density species and increases gas yields with both effects leading to a
reduction in liquid volume yields. Therefore, it is common practice to
charge the C.sub.6 paraffins to an isomerization unit to obtain C.sub.6
isoparaffin hydrocarbons. Consequently, octane upgrading commonly uses
isomerization to convert C.sub.6 and lighter boiling hydrocarbons and
reforming to convert C.sub.7 and higher boiling hydrocarbons.
The effluent from an isomerization zone will contain a mixture of more
highly branched and less highly branched paraffins. In order to further
increase the octane of the products from the isomerization zone, normal
paraffins, and sometimes less highly branched isoparaffins, are typically
recycled to the isomerization zone along with the feed stream in order to
increase the ratio of less highly branched paraffins to more highly
branched paraffins entering the isomerization zone. A variety of methods
are known to treat the effluent from the isomerization zone for the
recovery of normal paraffins and monomethyl branched isoparaffins for
recycling these less highly branched paraffins to the isomerization zone.
U.S. Pat. No. 2,966,528 issued to Haensel discloses a process for the
isomerization of C.sub.6 hydrocarbons and the adsorptive separation of
normal hydrocarbons from branched chain hydrocarbons. The process adsorbs
normal hydrocarbons from the effluent of the isomerization zone and
recovers the unadsorbed hydrocarbons as product, desorbs straight chain
hydrocarbons using a normal paraffin desorbent, and returns the desorbent
and adsorbed straight chain hydrocarbons to the isomerization zone.
U.S. Pat. No. 3,755,144 shows a process for the isomerization of a
pentane/hexane feed and the separation of normal paraffins from the
isomerization zone effluent. The isomerization zone effluent is separated
by a molecular sieve separation zone that includes facilities for the
recovery of desorbent from the normal paraffin containing stream that is
recycled to the isomerization zone. An extract stream that contains
isoparaffins is sent to a deisohexanizer column that separates isopentane
and dimethyl butane as a product stream and provides a recycle stream of
isohexane that is returned to the isomerization zone.
U.S. Pat. Nos. 4,717,784 and 4,804,802 disclose processes for the
isomerization of a hydrocarbon feed and the use of adsorptive separation
to generate normal paraffin and monomethyl-branched paraffin recycle
streams. The effluent from the isomerization zone enters a molecular sieve
separation zone that contains a 5A-type sieve and a ferrierite-type sieve
that adsorb normal paraffins and monomethyl-branched paraffins,
respectively. U.S. Pat. No. 4,804,802 discloses steam or hydrogen as the
desorbent for desorbing the normal paraffins and monomethyl-branched
paraffins from the adsorption section and teaches that steam or hydrogen
may be recycled with the normal paraffins or monomethyl-branched paraffins
to the isomerization zone.
One method of separating normal paraffins from isoparaffins uses adsorptive
separation under liquid phase conditions. In such methods, the
isomerization effluent contacts a solid adsorbent having a selectivity for
normal paraffins to effect the selective adsorption of normal paraffins
and allow recovery of the isoparaffins as a high octane product.
Contacting the normal paraffin containing adsorbent with the desorbent
material in a desorption step removes normal paraffins from the adsorbent
for recycle to the isomerization zone. Both the isoparaffin and normal
paraffin containing streams undergo a separation for the recovery of
desorbent before the isoparaffins are recovered as a product and the
normal paraffins recycled to the isomerization zone. Liquid phase
adsorption has been carried out in conventional swing bed systems as shown
in U.S. Pat. No. 2,966,528. The use of simulated moving bed systems for
the selective adsorption of normal paraffins is also known and disclosed
by U.S. Pat. No. 3,755,144. Simulated moving bed systems have the
advantage of increasing recovery and purity of the adsorbed and
non-adsorbed components in the isomerization zone effluent for a given
unit of adsorbent material.
In liquid phase adsorption systems the adsorbent contains selective pores
that will more strongly adsorb the selectively adsorbed components in the
feed mixture. The selective pore volume is limited and the quantity of
such pores must accommodate the desired volume of components to be
adsorbed from the feed mixture. The desorbent material is also a
selectively adsorbed component. Therefore, an extract column is typically
used to recover desorbent, otherwise any desorbent that passes through the
reactors of the reactors of the isomerization zone and enters the
adsorption section increases the amount of adsorbed component in the feed
mixture and requires additional adsorbent. If the quantity of selectively
adsorbed components is increased without increasing the available
selective pore volume for a given unit of feed, it was believed that the
purity of the extract and raffinate streams from the adsorption section
decreases. Therefore, the extract column has been viewed as necessary for
the desorption stage of the adsorption section since the loaded adsorbent
contains normal paraffins and desorbent material as adsorbed components
and all of these adsorbed components must be displaced by the desorbent.
Without the extract column, desorbent flow during the desorption step
would increase if traditional desorbent to pore volume ratios are
maintained thereby placing a greater quantity of desorbent in circulation
and increasing the amount of selective pore volume needed during the feed
step of the adsorption process. Under the conventional system, without
some method of rejecting desorbent material from the recycled extract
stream, the selective pore volume and desorbent requirements would
continue to progressively increase.
Most moving bed adsorption processes also use a desorbent material that has
a different composition than the primary components in the feed stream to
the adsorption section. As a result the desorbent material is typically
recovered from the raffinate material that it has desorbed for reuse in
the adsorption section. It has been the usual practice to use a raffinate
column to separate the desorbent material from the raffinate stream.
It is an object of this invention to simplify the adsorption section of a
combination adsorption and isomerization process.
It is a further object of this invention to increase the octane number of a
product stream that can be obtained from a combination of an isomerization
process and adsorptive separation section for the production of high
octane gasoline blending components.
It is a yet further object of this invention to make processes for the
isomerization of hydrocarbons and the liquid phase adsorptive separation
of isomerization effluents more economical.
Another object of this invention is to reduce the necessary equipment for
the liquid phase adsorptive separation of normal and isoparaffins.
Another object of this invention is to provide a more cost effective
arrangement for an isomerization of normal paraffins and the recycle of
normal paraffins using liquid phase adsorptive separations.
SUMMARY OF THE INVENTION
Applicants have discovered that the combination of an isomerization zone
for the isomerization of C.sub.5 -C.sub.6 paraffins and an adsorptive
separation section for the recycle of low octane paraffins to the
isomerization section can be operated with a single fractionation column
for the separation of raffinate, extract, product, desorbent and heavier
hydrocarbon components. In broad terms the invention is an arrangement for
a combination of an isomerization section for the isomerization of C.sub.5
and C.sub.6 paraffins and an adsorptive separation section for the
separation of the isomerization zone effluent. This arrangement is
structured such that the effluent from the isomerization zone enters an
adsorption section that separates the effluent into a raffinate stream and
an extract stream by contact of the effluent with an adsorbent and desorbs
the adsorbed components from the adsorbent using a desorbent material. The
raffinate from the adsorption section and a normal hexane stream enter a
deisohexanizer that supplies an overhead isomerate product stream, a
bottoms stream of heavy hydrocarbons and a sidecut stream of desorbent
material comprising normal hexane. Any excess desorbent and extract from
the adsorption section are recycled and combined with the feed entering
the isomerization zone. The direct recycle of extract from the adsorption
section to the isomerization zone, the transfer of the raffinate to the
deisohexanizer column and the recovery of desorbent from the
deisohexanizer eliminates the need for separate raffinate and extract
columns as are typically used in the prior art. It has also been
surprisingly found that this arrangement will increase the octane of the
isomerate product recovered overhead from the deisohexanizer. The octane
increase is the result of the recovery of monomethylpentanes with the
desorbent material as it is removed from the deisohexanizer column. The
monomethylpentanes are thereby recycled through the isomerization zone and
converted to higher octane isomers. As a result, this invention has the
advantage of simplifying the facilities for an isomerization adsorption
section combination while also increasing the octane number of the product
that is obtained therefrom.
Accordingly in a broad embodiment, this invention is a process for the
isomerization of a feedstream comprising C.sub.5 -C.sub.6 hydrocarbons.
This process separates the feedstream comprising C.sub.5 and C.sub.6
hydrocarbons into a first stream comprising normal C.sub.6 and higher
boiling hydrocarbons and a second stream comprising the lower boiling
components of the feedstream. A combination of the second stream and an
extract stream provides a combined feed which passes into an isomerization
zone containing an isomerization catalyst at isomerization conditions to
isomerize the hydrocarbons in the second stream and recover an
isomerization zone effluent from the isomerization zone that comprises
C.sub.5 and C.sub.6 isoparaffins. At least a portion of the effluent
passes from the isomerization zone to an adsorption section where it
contacts an adsorbent that selectively adsorbs normal paraffins from the
effluent and produces a raffinate stream having a decreased concentration
of at least one normal paraffin component relative to the effluent stream
and an adsorbent material having normal paraffins adsorbed thereon. At
least a portion of the raffinate stream and the first stream pass into a
fractionation column. Hydrocarbons having boiling points greater than
normal hexane are withdrawn from the bottom of the fractionation zone. An
isomerate product is withdrawn from the top of the fractionation zone. A
sidecut stream comprising normal hexane is withdrawn from the
fractionation zone. The desorbent stream passes to the adsorption zone
where it contacts an adsorbent material having normal paraffins adsorbed
thereon to desorb at least one normal paraffin component and produce the
extract stream.
In another embodiment, this invention is a process for the isomerization of
a C.sub.5 + naphtha feedstream comprising C.sub.5 -C.sub.6 hydrocarbons.
The process separates the feedstream in a splitter column into a first
stream comprising normal hexane and higher boiling hydrocarbons and a
second stream comprising a lower boiling component from the feedstream.
The combination of the second stream with an extract stream and a
desorbent stream produce a combined feed that passes into an isomerization
zone containing an isomerization catalyst at isomerization conditions to
isomerize the hydrocarbons in the combined feed. An isomerization zone
effluent comprising C.sub.5 and C.sub.6 isoparaffins is recovered from the
isomerization zone. C.sub.4 and lower boiling hydrocarbons are separated
from the effluent stream which is then passed to a simulated moving bed
adsorption section as an adsorber feed. Normal pentane is separated from
the adsorber feed by maintaining a net fluid flow through at least three
operationally distinct and serially interconnected zones of adsorbent in
the adsorption section. One zone is an adsorption zone defined by the
adsorbent located between a feed input stream at an upstream boundary of
the adsorption zone and a raffinate output stream at a downstream boundary
of the adsorption zone. Another zone is a purification zone defined by the
adsorbent located between an extract output stream at an upstream boundary
of the purification zone and the feed input stream at a downstream
boundary of the purification zone. And the third zone is a desorption zone
located immediately upstream from the purification zone that is defined by
the adsorbent located between a desorbent input stream at an upstream
boundary of the zone and the extract output stream at a downstream
boundary of the zone. The adsorber feed is passed into the adsorption zone
at adsorption conditions to effect the selective adsorption of the normal
pentane by the adsorbent in the adsorption zone and withdrawing a
raffinate output stream from the adsorption zone. At least a portion of
the desorbent stream is passed into the desorption zone at desorption
conditions to effect the displacement of normal pentane from the adsorbent
in the desorption zone. An extract output stream comprising normal
paraffins and desorbent is withdrawn from the desorption zone. A raffinate
output stream comprising isoparaffins and desorbent is withdrawn from the
adsorption zone. Periodically the feed input stream, raffinate output
stream, desorbent input stream and extract output stream input points are
advanced periodically through the column of adsorbent in a downstream
direction with respect to fluid flow in the adsorption zone to effect the
shifting of zones through the adsorbent and the production of extract
output and raffinate output streams. At least a portion of the raffinate
output stream is passed into an upper half of a fractionation column and
at least a portion of the first stream is passed into a lower half of the
fractionation column. A heavy hydrocarbon stream having a boiling point
greater than normal hexane is withdrawn from the bottom of the
fractionation column and an isomerate product stream is withdrawn from the
top of the fractionation column. The isomerate product stream is
essentially free of normal hexane and higher boiling hydrocarbons. A
desorbent stream is withdrawn as a sidecut from the fractionation column
at a tray location intermediate the column locations where the first
stream and raffinate stream enter the column.
Other objects, embodiments, and aspects of this invention are described in
the following detailed description of the invention.
BRIEF DESCRIPTION OF THE DRAWING
The figure is a schematic diagram of a flow arrangement for a combination
isomerization and adsorption process arranged in accordance with this
invention.
DETAILED DESCRIPTION OF THE INVENTION
This invention uses the combination of an isomerization zone and an
adsorptive separation section. The invention is not restriced to any
particular type of isomerization zone or adsorption section. The
isomerization zone can consist of any type of isomerization zone that
takes a stream of C.sub.5 -C.sub.6 straight chain hydrocarbons or a
mixture of straight chain and branched chain hydrocarbons and converts
straight chain hydrocarbons in the feed mixture to branched chain
hydrocarbons and branched hydrocarbons to more highly branched
hydrocarbons thereby producing an effluent having branched chain and
straight chain hydrocarbons. The adsorption sections is preferably liquid
phase and can utilize any type of well known adsorption process such as a
swing bed, simulated moving bed, or other schemes for contacting the
adsorbent with the feed mixture and desorbing the feed mixture from the
adsorbent with the desorbent material.
Suitable feedstocks for this process will include C.sub.5 and C.sub.6
hydrocarbons. At minimum the feed will include normal hexane and normal
pentane. The typical feed for this process will be a naphtha feed with an
initial boiling point in the range of normal butane. The feedstocks that
can be used in this invention include hydrocarbon fractions rich in
C.sub.4 -C.sub.6 normal paraffins. The term "rich" is defined as a stream
having more than 50% of the mentioned component. Preferred feedstocks are
substantially pure normal paraffin streams having from 4-6 carbon atoms or
a mixture of such substantially pure normal paraffins. It is also
preferred that the feed contain at least 10% and preferably at least 20%
normal pentanes. Another requirement of the feed is that it contain enough
normal hexane to supply the desorbent requirements of this invention.
Useful feedstocks include light natural gasoline, light straight-run
naphtha, gas oil condensates, light raffinates, light reformate, light
hydrocarbons, and straight-run distillates having distillation end points
of about 77.degree. C. (170.degree. F.) and containing substantial
quantities of C.sub.4 -C.sub.6 paraffins. The feed may also contain low
concentrations of unsaturated hydrocarbons and hydrocarbons having more
than 6 carbon atoms. The concentration of these materials should be
limited to 10 wt. % for unsaturated compounds and 20 wt. % for heavier
hydrocarbons in order to restrict hydrogen consumption in cracking
reactions. The feed in any normal paraffin recycle are combined and
typically enter the isomerization zone with a hydrogen recycle stream.
This application is described with reference to FIG. 1. Reference to the
specific arrangement for this invention is not meant to limit it to the
details disclosed therein. Furthermore, FIG. 1 is a schematic illustration
and does not show a number of details for the process arrangement such as
pumps, compressors, valves, stabilizers and recycle lines which are well
known to those skilled in the art.
The process begins by separating normal hexane and higher boiling
hydrocarbons in the feed from hydrocarbons boiling below normal hexane.
FIG. 1 shows an arrangement wherein a C.sub.5 + naphtha is charged by line
10 to a naphtha splitter column 12. The naphtha splitter separates
isohexane and lower boiling hydrocarbons from normal hexane and higher
boiling hydrocarbons. Normal hexane and higher boiling hydrocarbons are
taken by line 14 to a fractionation zone in the form of a deisohexanizer
column 16. A line 18 carries the isohexane and lower boiling hydrocarbons
overhead from naphtha splitter 12. Naphtha splitter column 12 is not an
essential part of this invention but will be used in most arrangements
when processing a combined normal pentane and normal hexane feed. The
splitter column is often used since a substantial quantity of the normal
hexane that is charged to the process must be available for withdrawal as
desorbent from the fractionation zone. Thus, it is important that a
substantial quantity of normal hexane be present in the deisohexanizer
column. The presence of the substantial amounts of normal hexane in the
deisohexanizer column is necessary to supply the desorbent needs as
hereinafter explained. The splitter column 12 will usually be present to
prevent all of the normal hexane from being charged first to the
isomerization zone which may convert too high a quantity of the normal
hexane to lower boiling isomers and leave an inadequate amount of normal
hexane available in the fractionation column for withdrawal as desorbent.
In addition the charging of large amounts of unconverted normal hexane
through the adsorption section may unnecessarily increase the flowrate
through the adsorption section. Of course, in certain situations where a
separate stream of normal hexane is available, the process may be operated
without a splitter.
The isohexane and lower boiling hydrocarbon stream carried by line 18 is
mixed with an extract stream from the adsorption section carried by a line
20. The extract stream can be taken directly from the adsorption section
and combined with the isomerization feed without intermediate separation.
The extract stream will contain normal hexane and lower boiling
hydrocarbons made up primarily of normal paraffins and monomethyl-branched
paraffins. In addition to the normal hexane, the other hydrocarbons in the
extract stream will usually be normal pentane and monomethylpentanes.
Therefore, all of the hydrocarbon components in the extract stream are
susceptible to octane improvement by further processing through the
isomerization zone. In some cases there will be an excess of desorbent
that is withdrawn from the fractionation column. This excess is also
combined with the feed to the isomerization zone. In the arrangement of
FIG. 1, excess desorbent is carried by line 22 and combined with the
extract and feed from lines 20 and 18, respectively, to form a combined
feed carried by line 24 to isomerization zone 26.
Hydrogen is admixed with the feed to the isomerization zone in an amount
that will provide a hydrogen to hydrocarbon molar ratio of from 0.01 to 10
in the effluent from the isomerization zone. Preferably, the hydrogen to
hydrocarbon ratio is in the range of 0.05 to 5. Although no net hydrogen
is consumed in the isomerization reaction, the isomerization zone will
have a net consumption of hydrogen often referred to as the stoichiometric
hydrogen requirement which is associated with a number of side reactions
that occur. These side reactions include saturation of olefins and
aromatics, cracking and disproportionation. For feeds having a high level
of unsaturates, satisfying the stoichiometric hydrogen will require a
higher hydrogen to hydrocarbon ratio for the feed at the inlet of the
isomerization zone. Hydrogen in excess of the stoichiometric amounts for
the side reactions is often maintained in the reaction zone to provide
stability and conversion by compensating for variation in feed stream
compositions that alter the stoichiometric hydrogen requirements. Higher
hydrogen to hydrocarbon ratios are often used to prolong catalyst life by
suppressing side reactions such as cracking and disproportionation. When
such side reactions occur, they can reduce conversion and lead to
formation of carbonaceous compounds, usually referred to as coke, that
foul the catalyst.
It has recently been found that the hydrogen to hydrocarbon ratio in
isomerization zones that use a chlorided platinum alumina catalyst can be
reduced significantly. In such cases, it is desirable to reduce the amount
of hydrocarbon that enters the isomerization zone such that the hydrogen
to hydrocarbon ratio of the effluent from the isomerization zone is less
than 0.05. Reduced hydrogen to hydrocarbon ratios have been used based on
the finding that the amount of hydrogen needed for suppressing coke
formation need not exceed dissolved hydrogen levels. The amount of
hydrogen in solution at the normal conditions of the isomerization zone
effluent are preferably in a ratio of from 0.02 to 0.01. The amount of
excess hydrogen over the stoichiometric requirement that is required for
good stability and conversion is in a ration of 0.01 to less than 0.05.
When the hydrogen to hydrocarbon ratio exceeds 0.05, it is not economically
desirable to operate the isomerization zone without the recycle of
hydrogen to the isomerization zone. Therefore, in such cases, recovery
facilities for hydrogen from the effluent will be provided as hereinafter
described. Hydrogen may be added to the feed mixture in any manner that
provides the necessary control for the addition of the hydrogen.
The hydrogen and hydrocarbon feed mixture is contacted in the reaction zone
with an isomerization catalyst. The catalyst composites that can be used
in the isomerization zone include traditional isomerization catalysts.
Such catalysts include high chloride catalyst on an alumina base
containing platinum, and crystalline aluminosilicates or crystalline
zeolites. Suitable catalyst compositions of this type will exhibit
selective and substantial isomerization activity under the operating
conditions of the process.
The preferred isomerization catalyst for this invention is a chlorided
platinum alumina catalyst. The aluminum is preferably an anhydrous
gamma-alumina with a high degree of purity. The catalyst may also contain
other platinum group metals. The term platinum group metals refers to
noble metals excluding silver and gold which are selected from the group
consisting of platinum, palladium, germanium, ruthenium, rhodium, osmium,
and iridium. These metals demonstrate differences in activity and
selectivity such that platinum has now been found to be the most suitable
for this process. The catalyst will contain from about 0.1 to 0.25 wt. %
of the platinum. Other platinum group metals may be present in a
concentration of from 0.1 to 0.25 wt. %. The platinum component may exist
within the final catalytic composite as an oxide or halide or as an
elemental metal. The presence of the platinum component in its reduced
state has been found most suitable for this process. The chloride
component termed in the art "a combined chloride" is present in an amount
from about 2 to about 10 wt. % based upon the dry support material. The
use of chloride in amounts greater than 5 wt. % have been found to be the
most beneficial for this process. The inorganic oxide preferably comprises
alumina and more preferably gamma-alumina, eta-alumina, and mixtures
thereof.
There are a variety of ways for preparing the catalytic composite and
incorporating the platinum metal and the chloride therein. The method that
has shown the best results in this invention prepares the catalyst by
impregnating the carrier material through contact with an aqueous solution
of a water-soluble decomposable compound of the platinum group metal. For
best results, the impregnation is carried out by dipping the carrier
material in a solution of chloroplatinic acid. Additional solutions that
may be used include ammonium chloroplatinate, bromoplatinic acid or
platinum dichloride. Use of the platinum chloride compound serves the dual
function of incorporating the platinum component and at least a minor
quantity of the chloride into the catalyst. Additional amounts of halogen
must be incorporated into the catalyst by the addition or formation of
aluminum chloride to or on the platinum-aluminum catalyst base. An
alternate method of increasing the halogen concentration in the final
catalyst composite is to use an aluminum hydrosol to form the aluminum
carrier material such that the carrier material also contains at least a
portion of the chloride. Halogen may also be added to the carrier material
by contacting the calcined carrier material with an aqueous solution of
the halogen acid such as hydrogen chloride.
It is generally known that high chlorided platinum-alumina catalysts of
this type are highly sensitive to sulfur and oxygen-containing compounds.
Therefore, the use of such catalysts requires that the feedstock be
relatively free of such compounds. A sulfur concentration no greater than
0.5 ppm is generally required. The presence of sulfur in the feedstock
serves to temporarily deactivate the catalyst by platinum poisoning.
Activity of the catalyst may be restored by hot hydrogen stripping of
sulfur from the catalyst composite or by lowering the sulfur concentration
in the incoming feed to below 0.5 ppm so that the hydrocarbon will desorb
the sulfur that has been adsorbed on the catalyst. Water can act to
permanently deactivate the catalyst by removing high activity chloride
from the catalyst and replacing it with inactive aluminum hydroxide.
Therefore, water, as well as oxygenates, in particular C.sub.1 -C.sub.5
oxygenates, that can decompose to form water, can only be tolerated in
very low concentrations. In general, this requires a limitation of
oxygenates in the feed to about 0.1 ppm or less. The feedstock may be
treated by any method that will remove water and sulfur compounds. Sulfur
may be removed from the feed stream by hydrotreating. A variety of
commercial dryers are available to remove water from the feed components.
Adsorption processes for the removal of sulfur and water from hydrocarbon
streams are also well known to those skilled in the art.
As a class, the crystalline aluminosilicate or crystalline zeolite
catalysts comprise crystalline zeolitic molecular sieves having an
apparent pore diameter large enough to adsorb neopentane. A silica alumina
molar ratio SiO.sub.2 :Al.sub.2 O.sub.3 of greater than 3; less than 60
and preferably between 15 and 30 is desirable. In preferred form, the
zeolite will contain an equivalent percent alkali metal cations and will
have those AlO.sub.4 -tetrahedra not associated with alkali metal cations;
either not associated with any metal cations or associated with divalent
or other polyvalent metal cations. Usually the molecular sieve is a
mordenite molecular sieve which is essentially in the acid form or is
converted to the acid form. Particularly preferred catalysts of this type
for isomerization are disclosed in detail in U.S. Pat. Nos. 3,442,794 and
3,836,597.
A preferred composition of zeolitic catalyst for use in the present
invention comprises a Group VIII noble metal, a hydrogen form crystalline
aluminosilicate, and a refractory inorganic oxide with the catalyst
composition having a surface area of at least 580 m.sup.2 /g. Significant
improvements in isomerization performance are realized when the surface
area of the catalytic composite is at or above 580 m.sup.2 /g. A Group
VIII metal is incorporated into the catalytic composite to supply a
hydrogenation/dehydrogenation function and the preferred Group VIII noble
metal is platinum. The Group VIII noble metal is present in an amount from
about 0.01 to 5% by weight of the composite and preferably in an amount of
at least 0.15% by weight but not over 0.35% by weight. The zeolitic
catalytic composite may also contain a catalytically effective amount of a
promoter metal such as tin, lead, germanium, cobalt, nickel, iron,
tungsten, chromium, molybdenum, bismuth, indium, gallium, cadmium, zinc,
uranium, copper, silver, gold, tantalum, or one or more of rare earth
metals and mixtures thereof. The hydrogen-formed silica alumina has either
a three-dimensional or channel pore structure crystal lattice framework.
The three-dimensional aluminosilicates include both synthetic and naturall
occurring silica aluminas such as faujasites, which include X-type,
Y-type, ultrastable-Y, and the like. L-type, omega-type, and mordenite are
examples of the channel pore structure crystalline aluminosilicates.
Mordenite, in either naturally occurring or synthetic form are preferred,
particularly with a silica to alumina ratio of at least 16:1. The hydrogen
form aluminosilicate may be present in an amount within the range of 50 to
about 99.5 wt. %, preferably within the range of 75 to about 95 wt. %, and
a refractory inorganic oxide may be present in an amount within the range
of from 25 to about 50 wt. %.
Operating conditions within the isomerization zone are selected to maximize
the production of isoalkane product from the feed components. Temperatures
within the reaction zone will usually range from about
40.degree.-320.degree. C. (100.degree.-600.degree.F.). Lower reaction
temperatures are generally preferred since they usually favor equilibrium
mixtures of isoalkanes versus normal alkanes. Lower temperatures are
particularly useful in processing feeds composed of C.sub.5 and C.sub.6
alkanes where the lower temperatures favor equilibrium mixtures having the
highest concentration of the most branched isoalkanes. When the feed
mixture is primarily C.sub.5 and C.sub.6 alkanes temperatures in the range
of from 60.degree. to 160.degree. C. are preferred. Higher reaction
temperatures increase catalyst activity and promote the isomerization of
C.sub.4 hydrocarbons. The reaction zone may be maintained over a wide
range of pressures. Pressure conditions in the isomerization of C.sub.4
-C.sub.6 paraffins range from 700 to 7000 Kpag. Preferred pressures for
this process are in the range of from 2000 to 3000 Kpag. The feed rate to
the reaction zone can also vary over a wide range. These conditions
include liquid hourly space velocities ranging from 0.5 to 12 hr..sup.-1,
however, space velocities between 1 and 6 hr..sup.-1 are preferred. The
isomerization zone will usually operate at a LHSV of about 1.5.
Operation of the reaction zone with the preferred chlorided
platinum-alumina catalyst also requires the presence of a small amount of
an organic chloride promoter. The organic chloride promoter serves to
maintain a high level of active chloride on the catalyst as low levels are
continuously stripped off the catalyst by the hydrocarbon feed. The
concentration of promoter in the reaction zone is maintained at from 30 to
300 ppm. The preferred promoter compound is carbon tetrachloride. Other
suitable promoter compounds include oxygen-free decomposable organic
chlorides such as propyldichloride, butylchloride, and chloroform to name
only a few of such compounds. The need to keep the reactants dry is
reinforced by the presence of the organic chloride compound which converts
to hydrogen chloride. As long as the process streams are kept dry, there
will be no adverse effect from the presence of hydrogen chloride.
The isomerization zone usually includes a two-reactor system with a first
stage reactor and a second stage reactor in the reaction zone. The
catalyst used in the process is distributed equally between the two
reactors. It is not necessary that the reaction be carried out in two
reactors but the use of two reactors confer several benefits on the
process. The use of two reactors and specialized valving allows partial
replacement of the catalyst system without taking the isomerization unit
off stream. For the short periods of time during which replacement of
catalyst may be necessary, the entire flow of reactants may be processed
through only one reaction vessel while catalyst is replaced in the other.
The use of two reaction zones also aids in maintaining lower catalyst
temperatures. This is accomplished by having any exothermic reaction such
as hydrogenation of unsaturates performed in a first reaction vessel with
the rest of the reaction carried out in a final reaction vessel at more
favorable temperature conditions.
The effluent from the reactors enters a stabilizer that removes light gases
and butane from the effluent (not shown). The amount of butane taken off
from the stabilizer will vary depending upon the amount of butane entering
the process. The stabilizer normally runs at a pressure of from 800 to
1700 Kpaa.
When the isomerization zone is operated with a high hydrogen to hydrocarbon
ratio, a separator is usually placed ahead of the stabilizer. A
hydrogen-rich recycle gas stream is recovered from the separator and
recycled for combination with the feed entering the isomerization zone.
When the isomerization zone operates with very low hydrogen to hydrocarbon
ratios the separator is not needed and the effluent from the isomerization
zone may enter the stabilizer directly.
The bottoms stream from the stabilizer provides an isomerization zone
effluent stream comprising C.sub.5 and higher boiling hydrocarbons that
include normal paraffins for recycle and isoparaffin products. The
chlorides which may be present in the reaction zone will usually pose no
problem for the sorbent in the adsorption zone. In normal operation, any
chlorides that are present in the effluent from the isomerization zone
will be removed in the overhead from the stabilizer. However, where the
isomerization zone or separators downstream from the isomerization are
subject to upsets, it may be desirable to provide a guard bed of some type
to treat the stabilizer bottoms and prevent any carryover of chloride
compounds into the adsorption section.
The isomerization effluent is taken by line 28 and enters the adsorption
section 30 where it is contacted with an adsorbent in an adsorption zone.
The adsorption section of this invention is operated to primarily remove
the normal pentane fraction from the effluent of the isomerization zone.
This process is especially suited for adsorption systems that use multiple
ports for supplying the process streams to the adsorbent and divide the
adsorbent into a plurality of zones for adsorbing normal paraffins,
recovering isoparaffins, purifying the adsorbent, and desorbing the normal
paraffins. A well-known process of this type is the simulated
countercurrent moving bed system for simulating moving bed countercurrent
flow systems. Such systems have a much greater separation efficiency than
fixed molecular sieve bed systems. In the moving bed or simulated moving
bed processes, the retention and displacement operations are continuously
taking place which allows both continuous production of an extract and a
raffinate stream and the continual use of feed and desorbent streams. One
preferred embodiment of this process utilizes what is known in the art as
the simulated moving bed countercurrent flow system. The operating
principles and sequence of such flow system are described in U.S. Pat. No.
2,985,589 incorporated herein by reference. In such a system it is the
progressive movement of multiple liquid access points down a molecular
sieve chamber that simulates the upward movement of molecular sieve
contained in the chamber.
A number of specially defined terms are used in describing the simulated
moving bed processes. The term "feed stream" indicates a stream in the
process through which feed material passes to the molecular sieve. A feed
material comprises one or more extract components and one or more
raffinate components. An "extract component" is a compound or type of
compound that is more selectively retained by the molecular sieve while a
"raffinate component" is a compound or type of compound that is less
selectively retained. In this process normal hydrocarbons from the feed
stream are extract components while feed stream branched chain and cyclic
hydrocarbons are raffinate components. The term "extract component" as
used herein refers to a more selectively retained compound such as normal
hydrocarbons in this process. The term "displacement fluid" "or desorbent"
shall mean generally a material capable of displacing an extract
component. The term "desorbent" or "desorbent input stream" indicates the
stream through which desorbent passes to the molecular sieve. The term
"raffinate output stream" means a stream through which most of the
raffinate components are removed from the molecular sieve. The composition
of the raffinate stream can vary from about 100% desorbent to essentially
100% raffinate components. The term "extract stream" or "extract output
stream" means a stream through which an extract material which has been
displaced by a desorbent is removed from the molecular sieve. The
composition of the extract stream can also vary from about 100% desorbent
to essentially 100% extract components.
The term "selective pore volume" of the molecular sieve is defined as the
volume of the molecular sieve which selectively retains extract components
from the feedstock. The term "non-selective void volume" of the molecular
sieve is the volume of the molecular sieve which does not selectively
retain extract components from the feedstock. This non-selective void
volume includes the cavities of the molecular sieve which are not capable
of retaining extract components and the interstitial void spaces between
molecular sieve particles. The selective pore volume and the non-selective
void volume are generally expressed in volumetric quantities and are of
importance in determining the proper flow rates of fluid required to be
passed into an operational zone for efficient operations to take place for
a given quantity of molecular sieve.
When molecular sieve "passes" into an operational zone (hereinafter defined
and described) its non-selective void volume together with its selective
pore volume carries fluid into that zone. The non-selective void volume is
utilized in determining the amount of fluid which should pass into the
same zone in a countercurrent direction to the molecular sieve to displace
the fluid present in the non-selective void volume. If the fluid flow rate
passing into a zone is smaller than the non-selective void volume rate of
molecular sieve material passing into that zone, there is a net
entrainment of liquid into the zone by the molecular sieve. Since this net
entrainment is a fluid present in non-selective void volume of the
molecular sieve, it, in most instances, comprises less selectively
retained feed components.
In the preferred simulated moving bed process only four of the access lines
are active at any one time: the feed input stream, displacement or
desorbent fluid inlet stream, raffinate outlet stream, and extract outlet
stream access lines. Coincident with this simulated upward movement of the
solid molecular sieve is the movement of the liquid occupying the void
volume of the packed bed of molecular sieve. So that countercurrent
contact is maintained, a liquid flow down the molecular sieve chamber may
be provided by a pump. As an active liquid access point moves through a
cycle, that is, from the top of the chamber to the bottom, the chamber
circulation pump moves liquid through different zones which require
different flow rates. A programmed flow controller may be provided to set
and regulate these flow rates.
The active liquid access points effectively divide the molecular sieve
chamber into separate zones, each of which has a different function. In
this embodiment of the process, it is generally necessary that three
separate operational zones be present in order for the process to take
place although in some instances an optional fourth zone may be used.
The retention or extract zone, zone 1, is defined as the molecular sieve
located between the feed inlet stream and the raffinate outlet stream. In
this zone, the feedstock contacts the molecular sieve, an extract
component is retained, and a raffinate stream is withdrawn. Since the
general flow through zone 1 is from the feed stream which passes into the
zone to the raffinate stream which passes out of the zone, the flow in
this zone is considered to be a downstream direction when proceeding from
the feed inlet to the raffinate outlet streams.
Immediately upstream with respect to fluid flow in zone 1 is the
purification zone, zone 2. The purification zone is defined as the
molecular sieve between the extract outlet stream and the feed inlet
stream. The basic operations taking place in zone 2 are the displacement
from the non-selective void volume of the molecular sieve of any raffinate
material carried into zone 2 by the shifting of molecular sieve into this
zone and the displacement of any raffinate material retained within the
selective pore volume of the molecular sieve. Purification is achieved by
passing a portion of extract stream material leaving zone 3 into zone 2 at
zone 2's upstream boundary to effect the displacement of raffinate
material. The flow of material in zone 2 is in a downstream direction from
the extract outlet stream to the feed inlet stream.
Immediately upstream of zone 2 with respect to the fluid flowing in zone 2
is the displacement or desorption zone, zone 3. The desorption zone is
defined as the molecular sieve between the desorption inlet and the
extract outlet stream. The function of the desorption zone is to allow a
desorbent which passes into this zone to displace the extract component
which was retained in the molecular sieve during a previous contact with
feed in zone 1 in a prior cycle of operation. The flow of fluid in zone 3
is essentially in the same direction as that of zones 1 and 2.
In some instances, an optional buffer zone, zone 4, may be utilized. This
zone, defined as the molecular sieve between the raffinate outlet stream
and the desorbent inlet stream, if used, is located immediately upstream
with respect to the fluid flow to zone 3. Zone 4 would be utilized to
conserve the amount of desorbent utilized in the desorption step since a
portion of the raffinate stream which is removed from zone 1 can be passed
into zone 4 to displace desorbent present in that zone out of the zone
into the desorption zone. Zone 4 will contain enough desorbent so that
raffinate material present in the raffinate stream passing out of zone 1
and into zone 4 can be prevented from passing into zone 3 thereby
contaminating extract stream removed from zone 3. In the instances in
which the fourth operational zone is not utilized, the raffinate stream
passed from zone 1 to zone 4 must be carefully monitored in order that the
flow directly from zone 1 to zone 3 can be stopped when there is an
appreciable quantity of raffinate material present in the raffinate stream
passing from zone 1 into zone 3 so that the extract outlet stream is not
contaminated.
A cyclic advancement of the input and output streams through the fixed bed
of molecular sieve can be accomplished by utilizing a manifold system in
which the valves in the manifold are operated in a sequential manner to
effect the shifting of the input and output streams thereby allowing a
flow of fluid with respect to solid molecular sieve in a countercurrent
manner. Another mode of operation which can effect the countercurrent flow
of solid molecular sieve with respect to fluid involves the use of a
rotating disc valve in which the input and output streams are connected to
the valve and the lines through which feed input, extract output,
displacement fluid input and raffinate output streams pass are advanced in
the same direction through the molecular sieve bed. Both the manifold
arrangement and disc valve are known in the art. Specifically rotary disc
valves which can be utilized in this operation can be found in U.S. Pat.
Nos. 3,040,777 and 3,422,848, incorporated herein by reference. Both of
the aforementioned patents disclose a rotary type connection valve in
which the suitable advancement of the various input and output streams
from fixed sources can be achieved without difficulty.
In many instances, one operational zone will contain a much larger quantity
of molecular sieve than some other operational zone. For instance, in some
operations, the buffer zone can contain a minor amount of molecular sieve
as compared to the molecular sieve required for the retention and
purification zones. It can also be seen that in instances in which
desorbent is used which can easily displace extract material from the
molecular sieve that a relatively small amount of molecular sieve will be
needed in a desorption zone as compared to the molecular sieve needed in
the retention zone or purification zone. Since it is not required that the
molecular sieve be located in a single column, the use of multiple
chambers or a series of columns is within the scope of the invention.
It is not necessary that all of the input or output streams be
simultaneously used, and in fact, in many instances some of the streams
can be shut off while others effect an input or output of material. The
apparatus which can be utilized to effect the process of this invention
can also contain a series of individual beds connected by connecting
conduits upon which are placed input or output taps to which the various
input or output streams can be attached and alternately and periodically
shifted to effect continuous operation. In some instances, the connecting
conduits can be connected to transfer taps which during the normal
operations do not function as a conduit through which material passes into
or out of the process.
Reference can be made to D. B. Broughton U.S. Pat. No. 2,985,589, and to a
paper entitled "Continuous Adsorptive Processing--A New Separation
Technique" by D. B. Broughton presented at the 34th Annual Meeting of the
Society of Chemical Engineers at Tokyo, Japan on Apr. 2, 1969, both
references incorporated herein by reference, for further explanation of
the simulated moving bed countercurrent process flow scheme.
Although both liquid and vapor phase operations can be used in many
adsorptive type separation processes, liquid-phase operation is preferred
for this process because of the lower temperature requirements and because
of the higher yields of extract product that can be obtained with
liquid-phase operation over those obtained with vapor-phase operation.
Extract conditions will, therefore, include a pressure sufficient to
maintain liquid phase. Desorption conditions will include the same range
of temperatures and pressures as used for extract conditions.
In the operation of this process, at least a portion of the raffinate
output stream will be passed directly to a fractionation zone. The
fractionation zone will typically be a single fractionation column, the
general design and operation of which is well known to the separation art.
In the Figure a line 32 passes the raffinate directly to deisohexanizer
16.
The fractionation zone serves a variety of purposes. It provides an
overhead product stream that contains a high concentration of isopentane
and dimethylbutanes. Typically, the research octane number of the product
stream will be between 91 and 94. The isomerate product stream also
contains low concentrations of normal pentane, normal hexane and
monomethylpentanes. These relatively lower octane hydrocarbons are reduced
by the operation of the adsorption section which preferentially adsorbs
normal pentane and directly recycles the normal pentanes to the
isomerization zone in the extract stream and the withdrawal of normal
hexane and monomethylpentanes in large amounts from the fractionation zone
for use as a desorbent material in the adsorption section which also
eventually is recycled to isomerization zone. The desorbent is preferably
removed as a sidecut from a single fractionation column. In FIG. 1, line
36 is shown as a sidecut stream from the deisohexanizer column 16.
Desorbent can be withdrawn from any point below the input point of line
32. Thus, the raffinate stream can be withdrawn from below the input point
for line 14. The Figure shows the preferred withdrawal point for sidecut
stream 36 which is between the input point for the raffinate stream
carried by line 32 and the input point for the normal hexane and higher
boiling hydrocarbon stream carried by line 14. In the operation of a
fractionation zone having the arrangement of deisohexanizer 16, normal
hexane drops down the column from the inlet of line 32 and rises up the
column from the inlet of line 14. The location of normal hexane input
points above and below the withdrawal point for sidecut 36 provides a
stream that is rich in normal hexane as well as closely boiling
monomethylpentanes that are carried over into the sidecut stream. The
withdrawal of the desorbent as a liquid sidecut from the deisohexanizer
has the advantage of disengaging desorbent from the raffinate stream at
very low utility cost.
Heavier hydrocarbons are withdrawn from the fractionation column as a heavy
hydrocarbon stream. For the single column deisohexanizer, this heavy
hydrocarbon stream is withdrawn by a line 38. Where a full boiling range
naphtha is used as the feed to the process, the heavy hydrocarbon feed
will comprise a C.sub.7 + naphtha. This bottoms stream will ordinarily be
used as the feed in a reforming zone.
Line 36 provides a desorbent for the adsorption section that is passed from
line 36 to the adsorption section by a line 40. Depending on the
conditions in the adsorption section and the isomerization zone, the
amount of the desorbent available through line 36 may exceed that needed
for the adsorption section. This excess desorbent is diverted from line 36
to the previously described line 22 and enters the isomerization zone
directly as part of the isomerization zone feed.
Excess desorbent in line 36 is present in the deisohexanizer as part of the
normal hexane that has been charged to the fractionation zone. Although it
may be possible to eliminate the excess desorbent from the fractionation
zone by changing the operation of a splitter column when one is provided,
it is usually desirable to have the excess desorbent in the fractionation
zone in order to improve the carryover of monomethylpentanes into the
isomerization zone.
In this invention, the extract stream does not enter any separation section
for the recovery of the displacement fluid. At least a portion of the
extract stream is recycled directly to the isomerization zone to provide
the recycle stream as previously described. The direct recycle of the
extract stream eliminates the need for a separation column and the
equipment associated therewith. The elimination of the separation column
for the extract stream significantly reduces the cost of the adsorption
section.
Prior art processes provided a column for the separation of desorbent from
the extract stream in the belief that the process could not be
economically operated without such a separation. In an adsorptive
separation process, the amount of potentially adsorbed component in the
feed that enters the adsorption zone will control the amount of selective
pore volume that must be available in the adsorbent and the amount of the
displacement fluid or desorbent that is needed to recover the adsorbed
material from the adsorbent. Looking more specifically at the process for
the separation of normal paraffins, the amount of normal paraffins in the
feed mixture sets the amount of selective pore volume that must be
available to process a given quantity of the feed mixture. In the case of
a simulated moving bed process, an excess of adsorbent to the amount of
normals in the feed mixture must be provided to adsorb all of the normals
in the feed. In order to fully desorb all of the adsorbed components from
the adsorbent, a large excess of displacement fluid or desorbent material
is also needed. The circulation of the selective pore volume at a rate
greater than the volumetric addition of normal paraffins and the
circulation of desorbent at a rate greater than the circulation of the
selective pore volume will not permit all of the desorbent material from
the desorption zone to reenter the extract zone with the feed material
unless there is a constantly increasing circulation of selective pore
volume. Therefore, some removal of desorbent material from the extract
stream is necessary in order to continuously operate the process with a
constant circulation of selective pore volume. The recycle of the extract
stream to the isomerization zone provides the necessary removal of
desorbent material from the process. The amount of normal hexane desorbent
present in the process is decreased by passing it through the
isomerization zone. A further control on the amount of desorbent that is
passed through the adsorption section is provided by diverting desorbent
from the fractionation zone directly to the isomerization zone as
previously described.
EXAMPLE
The ability of this isomerization zone and adsorption section combination
to operate without extract or raffinate columns and provide a high octane
isomerate product are demonstrated by the following example. This example
consists of engineering calculations that are based on experience from the
operation of similar components in commercial processing units. This
example is arranged in accordance with the isomerization zone and
adsorption section shown in the Figure and will be described using the
reference numbers appearing therein. A C.sub.5 + naphtha feed having the
composition given in Table 1 for stream No. 10 is fed into a naphtha
splitter at a temperature of 66.degree. C. and a pressure of 700 Kpaa.
Table 1 shows the flowrate of the various feed components into the naphtha
splitter. The naphtha splitter is arranged with 40 trays and operates with
a molar reflux to feed ratio of 0.6. A bottoms stream is taken from
splitter 12 by line 14 at a temperature of 120.degree. C. and a pressure
of 280 Kpaa and transferred to deisohexanizer column 16. An overhead
stream is taken by line 18 at a pressure of 250 Kpaa and a temperature of
60.degree. C. The overhead and bottoms stream have a flowing composition
as given in Table 1 under lines 18 and 14, respectively.
TABLE 1
__________________________________________________________________________
Flowing
Composition, Vol. %
Component 10 14 18 24 20 22 28 32 34 38 36
__________________________________________________________________________
i-Butane 0.06
0.00
0.35
0.10
0.01
0.00
0.32
0.27
0.60
0.00
0.00
n-Butane 0.77
0.00
4.20
1.40
0.42
0.00
0.57
0.25
0.56
0.00
0.00
i-Pentane 4.66
0.00
25.48
7.30
0.79
0.00
20.66
17.85
38.90
0.00
0.00
n-Pentane 8.81
0.00
48.18
18.86
8.92
0.00
6.42
0.29
0.62
0.00
0.00
n-Hexane 7.65
9.07
1.32
17.14
22.14
31.78
6.56
15.25
0.43
0.34
31.78
2-Me-Pentane
4.15
2.84
9.99
15.58
17.59
17.73
16.80
16.98
13.88
0.00
17.73
3-Me-Pentane
2.34
2.29
2.52
11.36
14.53
14.81
9.52
10.28
2.89
0.00
14.82
2,2-DiMe-Butane
0.09
0.01
0.44
0.87
1.08
0.41
17.93
15.55
33.21
0.00
0.42
2,3-DiMe-Butane
0.49
0.26
1.52
2.76
3.21
3.14
5.59
5.27
6.82
0.00
3.14
Cyclopentane
1.02
0.02
5.51
1.51
0.06
0.03
1.06
0.92
2.00
0.00
0.02
Cyclohexane
4.13
5.05
0.02
9.34
12.70
12.99
7.10
7.94
0.00
4.53
12.98
Me-Cyclopentane
3.79
4.59
0.24
11.25
15.20
15.66
6.09
7.44
0.10
0.33
15.63
Benzene 1.00
1.21
0.07
1.36
1.82
1.91
0.00
0.26
0.00
0.11
1.90
Methane 0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
Ethane 0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
0.00
Propane 0.03
0.00
0.15
0.04
0.00
0.00
0.00
0.00
0.00
0.00
0.00
C7+ 61.01
74.66
0.00
1.12
1.53
1.54
1.39
1.44
0.00
94.69
1.57
Total 100.00
100.00
100.00
100.00
100.00
100.00
100.00
100.00
100.00
100.00
100.00
__________________________________________________________________________
The feed components carried by line 18 are combined with an extract stream,
an excess desorbent stream carried by lines 20 and 22, respectively. The
flowing compositions of lines 20 and 22 are given in the Table along with
the flowing composition of a combined feed that enters an isomerization
zone via line 24 at a temperature of 83.degree. C. and a pressure of 250
Kpaa. In the isomerization zone, the combined feed is contacted with an
isomerization catalyst that comprises a chlorided platinum-alumina
catalyst at a liquid hourly space velocity of 1.5. As the combined feed
enters the isomerization zone, it is combined with hydrogen in an amount
to produce a hydrogen/hydrocarbon ratio of 0.05 at the outlet of the
isomerization section. The isomerization section includes a two-reactor
system that operates at a pressure of 3100 Kpaa. A stabilized effluent
having the composition given for line 28 is recovered from the
isomerization zone 26 and transferred as the feed to an adsorption section
30.
The adsorption section is arranged with an eight bed adsorption column
filled with a zeolite adsorbent of the Ca-A type. The adsorption section
operates at a cycle time of less than 60 minutes for a complete sequence
of all the zones through the beds of adsorbent. An operating temperature
of 93.degree. C. (200.degree. F.) and an operating pressure of 2760 Kpaa
are maintained within the adsorption section. The extract stream taken by
line 20 is withdrawn from the adsorption section along with a raffinate
stream having the composition given in the Table under line 32. The
raffinate stream is transferred without intermediate separation into the
deisohexanizer column 16.
Deisohexanizer column 16 is arranged with 100 trays and operates with a
molar reflux to net deisohexanizer overhead ratio of 4.0. The raffinate
stream enters the deisohexanizer at tray level 20. The desorbent stream
having a composition given under line 36 in the Table is withdrawn from
the deisohexanizer as a sidecut at tray level 65. All of the desorbent
taken by line 36 enters the adsorption section except for the amount of
excess desorbent which is carried by line 22 as previously described. The
previously described bottoms stream taken by line 14 enters the
deisohexanizer column at tray level 85. A bottoms stream having the
composition given for line 38 is withdrawn from the deisohexanizer column
and used as the feed to a reforming process. The contents of line 38 are
taken from the column at a temperature of 93.degree. C. and a pressure of
480 Kpaa. The remaining column output is taken overhead by line 34.
Line 34 recovers an isomerate product stream having the composition given
in Table 1. The contents of line 34 are recovered at a temperature of
38.degree. C. and a pressure of 350 Kpa. The isomerate product has the
properties given in Table 2.
TABLE 2
______________________________________
RONC 91.0
MONC 89.6
RVP 12.6
S.G. 0.6463
______________________________________
This example shows that the isomerate product has a high octane number, 3
to 5 octane numbers higher than that usually achievable with conventional
recycle isomerization schemes. Therefore, the flow arrangement of this
invention will improve the operation of an isomerization zone and
adsorption section combination by increasing the octane of the isomerate
obtained therefrom and simplifying the overall operation of the
combination process.
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